Pre-converter for methanol synthesis

ABSTRACT

A reactor system, plant and a process for the production of methanol from synthesis gas is described in which the reactor system comprises:
     (a) a first reactor adapted to be maintained under methanol synthesis conditions having inlet means for supply of synthesis gas and outlet means for recovery of a first methanol-containing stream, said first reactor being charged with a first volume of a methanol synthesis catalyst through which the synthesis gas flows and on which in use, partial conversion of the synthesis gas to a product gas mixture comprising methanol and un-reacted synthesis gas will occur adiabatically; and   (b) a second reactor adapted to be maintained under methanol synthesis conditions having inlet means for supply of the gaseous first methanol-containing stream, outlet means for recovery of a second methanol-containing stream and cooling means, said second reactor being charged with a second volume of a methanol synthesis catalyst through which the gaseous first methanol-containing stream flows and on which, in use, further conversion of the synthesis gas to a product gas mixture comprising methanol will occur.

This application is a divisional of U.S. patent application Ser. No.10/158,752, filed May 30, 2002.

BACKGROUND OF THE INVENTION

The present invention relates to a process, reactor system and plant forthe production of methanol. In particular, it relates to a process,reactor system and plant for producing methanol from hydrogen and carbonoxides.

Methanol is synthesised in large volumes annually by the conversion of acarbonaceous feedstock, such as natural gas, into a mixture of carbonoxides and hydrogen. Such a mixture of gases in often referred to assynthesis gas.

The conversion of a hydrocarbon-containing feedstock, such as naturalgas, into synthesis gas can be achieved by steam reforming, by partialoxidation, or by a combination of these processes.

In steam reforming a mixture of desulphurised hydrocarbon feedstock,such as natural gas, and steam is passed at high temperature, typicallyat a temperature of from about 600° C. to about 1000° C., and elevatedpressure, typically from about 10 bar up to about 50 bar, over asuitable reforming catalyst, such as a supported nickel catalyst. Onecommercially recommended catalyst which is suitable for this purposeuses a mixture of calcium and aluminum oxides as support for the nickel.The principal reaction is:CH₄+H₂O≈CO+H₂.

The reaction products themselves are further subject to the reversible“water gas shift” reaction in which carbon dioxide and hydrogen areproduced from carbon monoxide and steam:CO+H₂O≈CO₂+H₂.

Another method for producing synthesis gas involves the use, wholly orin part, depending upon the carbon to hydrogen ratio in thehydrocarbonaceous feedstock, of direct catalytic or non-catalyticpartial oxidation or secondary/autothermal reforming with oxygen. In thecase of methane this occurs according to the following equation:CH₄+½O₂≈CO+H₂.

A combination of steam reforming and partial oxidation orsecondary/autothermal reforming can also be used.

Conversion of the carbon oxides and hydrogen to methanol occursaccording to the following reactions:CO+2H₂≈CH₃OHCO₂+3H₂≈CH₃OH+H₂O.

These reactions are conventionally carried out by contacting thesynthesis gas with a suitable methanol synthesis catalyst under anelevated synthesis gas pressure, typically in the range of from about 50bar up to about 100 bar, usually about 80 bar, and at an elevatedmethanol synthesis temperature, typically from about 210° C. to about270° C. or higher, e.g. up to about 300° C.

Suitable methanol synthesis catalysts include copper containingcatalysts with a catalyst comprising a reduced zinc oxide/copper oxidemixture being particularly-suitable.

As with many reactions it is desirable to achieve the maximum rate ofreaction per weight of catalyst or per volume of the reactor.

A conventional methanol synthesis plant can be considered to comprisefour distinct parts, namely:

-   1. a reforming plant, which produces a mixture of carbon oxides and    hydrogen from a hydrocarbon feedstock;-   2. a compression stage which lifts the carbon oxides and hydrogen    mixture to a higher pressure suitable for downstream methanol    synthesis;-   3. a methanol synthesis section, in which crude methanol is produced    from the carbon oxides and hydrogen; and-   4. a distillation section, in which the final refined methanol    product is produced from the crude methanol.

A number of different types of reformer for use in part 1 of themethanol synthesis plant i.e. the reforming plant, are known in the art.One such type is known as a “compact reformer” and is described inWO-A-94/29013, which is incorporated herein by reference and whichdiscloses a compact endothermic reaction apparatus in which a pluralityof metallic reaction tubes are close-packed inside a reformer vessel.Fuel is burned inside the vessel, which comprises air and fueldistribution means to avoid excessive localised heating of the reactiontubes. In a compact reformer of this type heat is transferred from theflow gas vent and from the reformed gas vent of the reformer to incomingfeedstock, fuel and combustion air. Other types of reformer are not asefficient as the compact reformer in transferring heat internally inthis way. However, many other reformer designs are known and some aredescribed in EP-A-0033128, U.S. Pat. No. 3,531,263, U.S. Pat. No.3,215,502, U.S. Pat. No. 3,909,299, U.S. Pat. No. 4,098,588, U.S. Pat.No. 4,692,306, U.S. Pat. No. 4,861,348, U.S. Pat. No. 4,849,187, U.S.Pat. No. 49,090,808, U.S. Pat. No. 4,423,022, U.S. Pat. No. 5,106,590and U.S. Pat. No. 5,264,008, U.S. Pat. No. 5,264,008 and WO 98/28071which are incorporated herein by reference.

In a conventional plant, synthesis gas is compressed in passage from thereforming plant to the methanol synthesis zone. This compression stageis generally present in order to provide the required pressure of from50 bar to 100 bar in the methanol synthesis zone. The compressed gas isthen passed to the methanol synthesis section.

In U.S. Pat. No. 4,594,227 apparatus for carrying out a catalyticchemical reaction is described which comprises a vertical, annular,intercylinder space which is divided by radially extending verticalpartition walls into a plurality of chambers some of which includeheat-exchanging tubes. Segments containing no heat-exchanging tubes maybe packed with catalyst and utilised adiabatically to preheat thereaction gases. In use, the reaction gases will pass outwardly throughthis first segment, where any reaction will cause heating, they then aretransmitted through the annular space surrounding the intercylinderspace before travelling inwardly through the segment containing catalystand cooling tubes where further reaction will occur.

Whilst this arrangement does offer certain advantages, it also suffersfrom various drawbacks. A principle disadvantage arises from themulti-segmental radial flow. This flow pattern causes the gas velocityto vary as the flow traverses from the centre of the reactor to theoutside and back, due to the changing cross-sectional area of thesegments. This changing velocities of the segments causes the heattransfer coefficient between the reacting gases and the cooling mediumin the tubes to vary. In particular the heat transfer will increase asthe gas velocity increases and will decrease as the gas velocity isreduced.

Thus the multi-segmental arrangement of the radial flow apparatus inU.S. Pat. No. 4,594,227 does not allow the gas velocity pattern andresultant heat transfer pattern to be optimised.

Various methanol production processes are known in the art, andreference may be made, for example, to U.S. Pat. No. 5,610,202, U.S.Pat. No. 4,968,722, U.S. Pat. No. 5,472,986, U.S. Pat. No. 4,181,675,U.S. Pat. No. 5,063,250, U.S. Pat. No. 4,529,738, U.S. Pat. No.4,595,701, U.S. Pat. No. 5,063,250, U.S. Pat. No. 5,523,326, U.S. Pat.No. 3,186,145, U.S. Pat. No. 344,002, U.S. Pat. No. 3,598,527, U.S. Pat.No. 3,940,428, U.S. Pat. No. 3,950,369, WO-A-98/28248 and U.S. Pat. No.4,051,300 which are incorporated herein by reference.

Various suggestions have been made for modifications to the plant designwith a view to improving the economics of the production process.

Several suggestions for improving the efficiency of the reaction havebeen made which incorporate the use of multiple reaction stages. Forexample, in U.S. Pat. No. 5,631,302 it is suggested that the methanolsynthesis section should include two separate synthesis reactors. Inthis arrangement, the synthesis gas is passed to the first synthesisreactor, which is a shaft reactor containing a fixed bed of acopper-containing catalyst. The reaction in this shaft reactor iscarried out adiabatically and in the absence of any recycling ofsynthesis gas. The product stream from this first reactor, whichcontains methanol vapour, is cooled to condense the methanol which isseparated from the unreacted gaseous. components of the first productstream. These unreacted gaseous components are then compressed, heatedand fed to the second reactor where they react to form methanol. Thesecond reactor is preferably a tubular reactor in which the coppercatalyst is indirectly cooled by water which is boiling under highpressure. The product stream from the second reactor is cooled and themethanol is removed by separation. Any unreacted gaseous components arecompressed and heated before being returned to the second reactor.

Thus in U.S. Pat. No. 5,631,302 the first reactor is located outside themain reactor loop and simply serves to modify the composition of thefeed gas before it enters the main reaction loop. The arrangement ofU.S. Pat. No. 5,631,302 is said to be useful where the synthesis gasfeed has a CO₂:CO ratio which exceeds 2:1.

An alternative arrangement is suggested in U.S. Pat. No. 5,827,901. Inthis arrangement two synthesis reactors are connected in series suchthat the product stream from the first reactor is passed directly to theinlet of the second reactor. The first reactor is a water cooled reactorin which the catalyst is located in tubes through which the gaseousreactants flow. The second reactor may be selected from a variety ofdesigns. Whichever design is used, cooling in the second reactor isprovided by counter-current heat exchange with the feed synthesis gasbefore it is fed to the first reactor.

This arrangement allows for cooler exit temperatures from the secondreactor to be achieved than are conventionally achievable. However,whilst the lower temperature may allow the reaction equilibrium to movetowards completion, it may also reduce the rate of reaction, andtherefore may require more catalyst per unit of product.

Other examples include U.S. Pat. No. 5,427,760 in which two reactionstages are used in an attempt to achieve a higher overall conversion tothe desired ammonia than can be achieved in a single stage and U.S. Pat.No. 4,867,959 in which two or more reaction stages are described, withcooling between each stage, to increase conversion. As discussed byKobayashi and Green in a paper presented to the 1990 World MethanolConference, this approach can be extended to include a large number ofstages. This paper also illustrates the optimum rate line for methanolsynthesis.

Whilst an optimum rate line is known, a near-optimum reaction profile isnot practical in commercial arrangements. This is because such a profilewould generally require the reaction to start at high temperature andgradually fall as the reaction proceeds. Some suggestions have been madeto produce a system which approaches the optimum rate line, such asthose in the Kobayashi and Green paper however, a commerciallysatisfactory arrangement has not been realised.

Thus it will be understood that whilst the systems of the prior art gosome way to addressing the problems associated with reducing theoperating and/or investment costs of producing methanol, variousdisadvantages and drawbacks remain and there is still a requirement foralternative arrangements which will address at least some of theseproblems.

SUMMARY OF THE INVENTION

One alternative arrangement which goes at least some way to addressingthese problems is an arrangement in which the catalyst for theproduction of methanol in the methanol synthesis section of the plant isdivided into two volumes of independent geometry. An arrangement of thistype will allow the highest temperature to which the catalyst issubjected to be reduced whilst also reducing the overall reactorpressure drop.

Thus, according to a first aspect of the present invention there isprovided a reactor system for use in the production of methanol fromsynthesis gas comprising:

-   (a) a first reactor adapted to be maintained under methanol    synthesis conditions having inlet means for supply of synthesis gas    and outlet means for recovery of a first methanol-containing stream,    said first reactor being charged with a first volume of a methanol    synthesis catalyst through which the synthesis gas flows and on    which in use, partial conversion of the synthesis gas to a product    gas mixture comprising methanol and unreacted synthesis gas will    occur adiabatically; and-   (b) a second reactor adapted to be maintained under methanol    synthesis conditions having inlet means for supply of the gaseous    first methanol-containing stream, outlet means for recovery of a    second methanol-containing stream and cooling means, said second    reactor being charged with a second volume of a methanol synthesis    catalyst through which the gaseous first methanol-containing stream    flows outwardly from the inlet means and on which, in use, further    conversion of the synthesis gas to a product gas mixture comprising    methanol will occur. In one preferred arrangement the cooling means    is arranged such that heat transfer decreases as the gas flows from    the inlet to the outlet.

Thus the present invention provides an arrangement in which an adiabaticbed is combined with a downstream cooled catalyst bed. This arrangementallows a commercially acceptable process to be provided which, in themost preferred arrangement has a reaction profile which willapproximately follow the optimum rate line.

In a preferred arrangement, the first and second reactors are separatereactors and the outlet means of the first reactor is connected to theinlet means of the second reactor by conventional means.

However, the first and second reactors may be zones located within asingle reactor. In this latter arrangement it is important that thecatalyst of the first and second reactors do not have a common face andthe reactor system will therefore preferably include means forseparating the two catalyst volumes and for transferring the gaseousfirst methanol-containing stream from the outlet means of the firstreactor to the inlet means of the second reactor. It will be understoodthat in this arrangement the “outlet means” and “inlet means” may beareas of the reactor rather than specific items of construction. The oreach reactor is preferably a pressure vessel.

The first volume of the methanol synthesis catalyst is preferablyarranged as a horizontal volume and the apparatus is preferably arrangedsuch that the synthesis gas preferably flows through the catalyst volumein a vertical direction. In a particularly preferred. arrangement, thedepth of the first volume of catalyst is preferably less than itshorizontal dimensions. Thus, where the first volume of catalyst iscylindrical, the depth of the cylinder is less than the diameter of thecross-section of the cylinder or where it is a prism, the depth will beless than the cross-sectional dimensions.

One benefit associated with reducing the depth of the bed is that thepressure drop of the gases as they pass through the bed is reduced. Thishas the effect of reducing the overall cost of the plant. A furtheradvantage of the arrangement of the present invention is that theincreased cross-sectional area of the catalyst in the first volumepresented to the synthesis gas relative to the depth of the first volumeof catalyst through which the gas has to flow when compared to prior artcatalyst volumes in vertical vessels facilitates the heat transfer byconduction and/or radiation from the relatively hotter catalyst at thebottom of the bed to the relatively cooler catalyst located towards theupper surface of the bed. The resultant increase in the averagetemperature of the catalyst bed serves to increase the rate of reaction.

Whilst the horizontal arrangement of the catalyst volume is preferred,any suitable arrangement may be utilised provided that it is of a lowpressure drop design. In one alternative arrangement the first volume ofcatalyst in the first reactor allows for radial flow.

The first volume of catalyst will be retained in position within thereactor by any suitable means and is preferably located on a supportmeans, such as a grid, which allows the gaseous reactants to passthrough the catalyst volume with minimal reduction in gas pressure. Thefirst catalyst volume is preferably a fixed bed arrangement.

The first reactor may additionally include an inlet gas distributor toassist in achieving good distribution of the synthesis gas throughout anupper area of the reactor before the gas comes into contact with thecatalyst volume.

Any suitable arrangement for the second reactor may be used.Further.-any suitable arrangement for the second volume of catalyst inthe second reactor may be used. In a preferred arrangement, the reactoris designed for a minimum pressure drop. In a most preferredarrangement, the second volume of catalyst in addition to providingminimum pressure drop is also arranged to allow good heat transfer fromthe catalyst to a cooling means. The presence of the cooling means isparticularly preferred as it provides that the exit temperature of thegas is reduced towards the equilibrium value and is prevented fromrising significantly which would result in faster rates of catalystdeactivation.

In a particularly preferred arrangement, the product stream from thefirst reactor flows radially from a central inlet to an outlet collectorlocated at a reactor zone wall through the second catalyst volume.

The catalyst may be cooled by any suitable means. In one preferredarrangement, cooling is provided by boiling water cooling in tubes whichpass through the catalyst bed in a conventional manner. This method ofcooling allows steam to be produced which may then be used to drive thecompressor which may be present to increases the pressure of the feed orrecycle synthesis gas prior to its addition to the first reactor zone.

Producing the steam for use in driving the compressor in this manner hasvarious benefits. In particular, the overall efficiency of the plant maybe improved which will reduce the overall costs.

The catalyst in the first and second volumes may be the same ordifferent. The catalyst for use in the methanol synthesis in eachreactor is preferably selected from, but is not limited to,copper-containing catalysts, for example reduced CuO—ZnO catalysts.Preferred catalysts include those sold under the designation 51/8 by ICIKatalco. Other suitable catalysts are described in U.S. Pat. No.6,054,497 which is incorporated herein by reference.

According to a second aspect of the present invention there is provideda plant for the production of methanol from a synthesis gas mixturecomprising carbon oxides, hydrogen and methane comprising:

-   (a) a methanol synthesis zone including the reactor system according    to the above-mentioned first aspect of the present invention; and-   (b) a methanol recovery zone, adapted to be maintained under    methanol recovery conditions, for recovery of a crude methanol    product stream from the product gas mixture, and for recovery of a    vaporous stream comprising unreacted material of the synthesis gas.

In a preferred arrangement the plant additionally includes:

-   (c) means for recycling at least a portion of the unreacted material    of the synthesis gas from the methanol recovery system to the    methanol synthesis zone.

The synthesis gas mixture is preferably produced from a hydrocarbonfeedstock material in plant comprising a steam reforming zone, adaptedto be maintained under steam reforming conditions and charged with acatalyst effective for catalysis of at least one steam reformingreaction, for steam reforming of a vaporous mixture of the hydrocarbonfeedstock in the steam to form a synthesis gas mixture comprising carbonoxides, hydrogen and methane. Suitable steam reformers include thosedetailed above which are incorporated herein by reference.

The plant of the present invention may include a plurality of reactorsystems according to the above first aspect of the present invention.These may be located in parallel such that the overall plant productionof methanol may be increased or in an alternative arrangement, they maybe located in series such that the second methanol-containing stream ispassed either directly or indirectly to a first reactor zone of a secondreactor system. One benefit of this arrangement is that the recovery ofreactants is improved.

According to a third aspect of the present invention there is provided aprocess for producing methanol from a synthesis gas comprising:

-   (a) supplying the synthesis gas mixture to the methanol synthesis    reactor system of the above-mentioned first aspect of the present    invention maintained under methanol synthesis conditions;-   (b) recovering from the methanol synthesis reactor system a product    gas mixture comprising methanol and any unreacted material of the    synthesis gas mixture;-   (c) supplying material of the product gas mixture to a methanol    recovery zone maintained under methanol recovery conditions; and-   (d) recovering from the methanol recovery zone a crude methanol    product stream and a vaporous stream comprising unreacted material    of the synthesis gas mixture.

In a preferred arrangement, the process additionally includes the stepof recycling the unreacted material of the synthesis gas mixture to themethanol synthesis reactor.

The synthesis gas is preferably formed from a hydrocarbon feedstock in aprocess comprising contacting a vaporous mixture comprising thefeedstock and steam in the steam reforming zone with a catalysteffective for catalysis of at least one reforming reaction andrecovering from the reforming zone a synthesis gas mixture comprisingcarbon oxide, hydrogen and methane.

The synthesis gas is preferably compressed before being supplied to themethanol synthesis reactor system. The pressure of the gaseous reactantsentering the first reactor zone will preferably in the region of 20 barto 200 bar. The first volume of catalyst in the first reactor ispreferably arranged such that the gas pressure drop that occurs is lessthan 0.5 bar. The compression may occur by any suitable means. Typicallythe motive force of gas compression is provided by high pressure steamgenerated within the plant by a steam turbine. However, as has beendiscussed, the steam may be wholly or in part provided by the coolingsystem in the second reactor.

The temperature of the gaseous reactants entering the first reactor willpreferably be in the region of about 180° C. to about 220° C. Thereactants exiting the first reactor and entering the second reactor willbe substantially the same temperature. The temperature of these streamsis most preferably just below peak reaction temperature and willtherefore most preferably be in the region of about 230° C. to about350° C.

The space velocity of the synthesis gas mixture entering the firstvolume of catalyst is preferably in the region of 5 to 20% of the totalspace velocity, dependent on the syntheses gas composition.

The apparatus, plant and process of the present invention havesignificant advantages over conventional apparatus, plant and processesfor the production of methanol.

Most significantly the invention allows the first volume of catalyst tobe designed to give a low pressure drop, while the second volume ofcatalyst is designed to meet the additional requirement to controlreaction temperature. By separating the two volumes flexibility isattained which leads to enhanced performance of the reactor in terms ofreactor pressure drop, steam production pressure and reactor conversion.

One further benefit of the present invention is that the arrangement inthe first reactor zone is simple to manufacture which substantiallyreduces the cost of construction.

Whilst generally there is no economic benefit in dividing volumes ofmaterial in process plants, since this inevitably leads to an increasein aspects of reactor construction cost per unit volume, in the presentinvention substantial benefits are obtained.

For example, by reducing the reaction system pressure drop, therequirement to compress the recycle gas stream is reduced and hence thecost of recycle gas compression is reduced. A slight increase inconversion of synthesis gas to methanol may also be noted. However, moreimportantly, the improved efficiency of the present invention may meanthat the total catalyst volume required for a given methanol productionrate is reduced.

The system has the further benefit in that the maximum temperatureachieved in the reaction zones is reduced which will reduce the rate atwhich catalyst deactivation occurs.

BRIEF DESCRIPTION OF THE FIGURES

The present invention will now be described, by way of example, withreference to the accompanying drawings in which:

FIG. 1 is a representation of a reactor system in accordance with thepresent invention;

FIG. 2 is a schematic diagram of a process in accordance with thepresent invention;

FIG. 3 is a representation of a reactor system according to thecomparative example; and

FIG. 4 is a graph comparing results obtained in reactions in the systemof FIGS. 1 and 3.

DETAILED DESCRIPTION OF THE INVENTION

It will be understood by those skilled in the art that the drawings arediagrammatic and that further items of equipment such as feedstockdrums, pumps, vacuum pumps, compressors, gas recycling compressors,temperature sensors, pressure sensors, pressure relief valves, controlvalves, flow controllers, level controllers, holding tanks, storagetanks and the like may be required in a commercial plant. Provision ofsuch ancillary equipment forms no part of the present invention and isin accordance with conventional chemical engineering practice.

Referring to FIG. 1, the methanol reaction system of the presentinvention comprises a first reactor vessel 1 have a gas inlet 2 andthrough which in use synthesis gas at a temperature of 180° C. to 220°C. and a pressure of 20 bar to 200 bar enters. The gas is preferableevenly distributed within the upper area 3 of the vessel 1 by means ofan inlet gas distributor 4.

The synthesis gas then passes through a first volume of catalystcomprising a bed of a suitable methanol synthesis catalyst 5 supportedon a grid 6. The catalyst is preferably a copper-containing catalystsuch as CuO—ZnO. The bed of catalyst 5 has a high cross-section to depthratio when compared with conventional systems. The depth of the bed ispreferably between about 0.4 to about 1.2 metres. This low depth to thecatalyst bed reduces the pressure drop of the gaseous mixture as itpasses through the bed.

Some of the synthesis gas will undergo conversion adiabatically tomethanol as it is passed through the catalyst bed 5 such that the gasescollected by the gas collector 7 include un-reacted synthesis gas andmethanol. This then leaves reactor 1 via outlet 8, and the gaseousstream is passed in line 9 to the inlet 10 to the second reactor zone11.

This second reactor may be of any suitable design but preferably has acentral gas distributor 12 and is constructed to allow for radial gasflow from the central gas distributor 12 to the outlet collector 13 atthe vessel wall of reactor 11. Thus gas flow is predominantly radialsuch that the gaseous mixture will pass through the catalyst bed 14.

The bed 14 is cooled by water boiling in a plurality of tubes 33 whichpass through the catalyst bed between tube sheets 34 located above andbelow-the catalyst bed. Pressurised cooling water is introduced viainlet nozzle 35 and steam and water exits via the exit nozzle 36. Thereactor also includes a utility nozzle 37.

FIG. 2 illustrates the flow sheet of the process of the presentinvention in which synthesis gas, which may have been formed in a steamreforming plant (not shown) is fed at a pressure of from 20 bar to 200bar at a temperature of from about 180° C. to about 220° C. to the firstreactor volume 1 in stream 38 where reaction occurs adiabatically.

The first product stream which will be at a pressure from about 20 barto about 200 bar and a temperature of from about 200° C. to about 250°C. and which includes un-reacted synthesis gas and methanol is passed tothe second reactor volume 11 where further reaction occurs. The productstream 39 collected from the second reactor 11 is passed to the coolingheat exchanger train 40 where the methanol is condensed. Stream 41 istherefore of mixed phase including un-reacted synthesis gas andcondensed methanol. These are separated in separator 42 and a crudemethanol product stream is retrieved in line 43 for further purificationby conventional means such as distillation.

The gas stream 44 from the separator will generally be divided into arecycle stream 45 and a purge gas stream 51. The purge gas stream willremove inert materials and optionally excess hydrogen.

Gas stream 45 is then compressed in gas compressor 46 to a sufficientpressure to allow recycling to the first reactor zone 1. The resultantstream 47 may be combined with fresh synthesis gas via line 48. In onealternative arrangement, the fresh synthesis gas may be added into line45, i.e. before the recycle stream is passed through the compressor.

The resultant stream will then pass through heat exchangers 49 beforebeing passed in stream 50 to the first reactor zone 1.

In one arrangement the heat exchangers 40 and 49 may be combined suchthat the hot product stream 39 is cooled in counter-current heatexchange with the recycle stream 47 which is consequently warmed.

FIG. 3 illustrate an reactor of generally conventional design which willproduce the same quantity of methanol from the same gas feed as thatillustrated in FIG. 1. In this arrangement the conventional design hasbeen slightly modified to more closely be comparable to the arrangementof FIG. 1 and thus on an inner part of the catalyst volume 52 has nocooling tubes passing through and therefore an essentially adiabatic bedis achieved through which the synthesis gas flows after entry throughinlet 54.

As the gas flow is radial towards outlet 55, after passing through theinner volume, the gas passes through an other part 53 of the catalystvolume through which cooling tubes 56 pass. Thus the second part of thevolume functionally corresponds with the second reactor of FIG. 1.However, in order to achieve the same production rate while maintainingthe same vessel diameter, the height of the vessel has to be increasedover that required in the arrangement of FIG. 1. This is partly due tothe inclusion of the adiabatic volume of catalyst. However, it is alsodue to reduced gas velocities caused by the increased height over theincreased mean diameter of the cooled catalyst volume which causes alower heat transfer coefficient. Thus to remove the same amount of heatwithout increasing the peak catalyst temperature, a larger surface areaand hence a large volume of cooling tubes has to be provided. The lowergas velocities also result in poorer heat transfer and thus poorertemperature control within the cooled bed.

Further an increased volume of catalyst is required since the exittemperature of the gas from the catalyst bed is reduced which reducesthe rate of reaction per unit volume of catalyst. Thus, this arrangementwould be wholly uneconomical.

EXAMPLE 1 AND COMPARATIVE EXAMPLE 1

Reactive systems in accordance with those illustrated in FIG. 1 and FIG.3 are fed with a feed gas having the following composition:

Component Volume % Steam 0.4 Hydrogen 72.5 Carbon Monoxide 15.4 CarbonDioxide 7.3 Methane 3.3 Nitrogen 1.1 Total 100.0

The results obtained are as follows:

Comparative Example 1 Example 1 Catalyst Volume, relative 1.0 1.1Methanol Production Per Unit 1.0 0.9 Volume (relative) Generated Steampressure bar 25.5 25.5 Reactor Pressure Drop bar 1.4 1.15 Heat Transfersurface, relative 1.0 1.09 Inlet gas temperature ° C. 221 218

This performance data is for new catalyst which has been un-aged byplant operation. The distribution of catalyst volume verses temperaturewithin the reactors is illustrated in FIG. 4. As can be seen, from thearrangement of the present invention of Example 1 has 49% of thecatalyst operating above a temperature of 250° C. In contrast, in thesingle volume arrangement of Comparative Example 1 and FIG. 3, the totalamount of catalyst is increased to 110% of that of Example 1 and 60% ofthis catalyst is operating above 250° C.

Thus it would be understood that the arrangement of the presentinvention allows the amount the catalyst required to be reduced, for animproved performance with the same amount of the catalyst, and for thecatalyst to be operated at lower temperatures which will prolongcatalyst life.

Whilst the present invention has been discussed with regard to theproduction of methanol from synthesis gas; it will be understood thatthe reactor design of the present invention may also be applied to otherexothermic chemical reactions, such as the formation of ammonia.

1. A process for producing methanol from a synthesis gas comprising: (a)supplying the synthesis gas mixture to a methanol synthesis reactorsystem comprising: (i) a first reactor adapted to be maintained undermethanol synthesis conditions having an inlet for supply of synthesisgas and an outlet for recovery of a first methanol-containing stream,said first reactor being charged with a first volume of a methanolsynthesis catalyst through which the synthesis gas flows and on which inuse, partial conversion of the synthesis gas to a product gas mixturecomprising methanol and un-reacted synthesis gas will occuradiabatically, and (ii) a second reactor adapted to be maintained undermethanol synthesis conditions having an inlet for supply of the gaseousfirst methanol-containing stream which has not been subjected to coolingafter being recovered from the outlet of the first reactor, an outletfor recovery of a second methanol-containing stream and indirectcoolings means, said second reactor being charged with a second volumeof a methanol synthesis catalyst through which the gaseous firstmethanol-containing, a stream flows outwardly from the inlet and onwhich, in use, further conversion of the synthesis gas to a product gasmixture comprising methanol will occur, wherein, when said first andsecond reactors are zones in a single reactor, the first catalyst andthe second catalyst are separated from each other wherein said system ismaintained under methanol synthesis conditions; (b) recovering from themethanol synthesis reactor system a product gas mixture comprisingmethanol and an un-reacted material of the synthesis gas mixture; (c)supplying material of the product gas mixture to a methanol recoveryzone maintained wider methanol recovery conditions; and (d) recoveringfrom the methanol recovery zone a crude methanol product stream and avaporous stream comprising un-reacted material of the synthesis gasmixture.
 2. A process according to claim 1 additionally including thestep of recycling the un-reacted material to the methanol synthesisreactor.
 3. The process according to claim 1 wherein the synthesis gasis formed from a hydrocarbon feedstock in a process comprisingcontacting a vaporous mixture comprising the feedstock and steam in thesteam reforming zone with a catalyst effective for a catalysis of atleast one reforming reaction and recovering from the reforming zone asynthesis gas mixture comprising carbon oxide, hydrogen and methane. 4.The process according to claim 1 wherein the synthesis gas is compressedbefore being supplied to the methanol synthesis reactor system.
 5. Theprocess according to claim 1 wherein the pressure of the gaseousreactants entering the first reactor zone are in the region or 20bar to200 bar.
 6. The process according to claim 1 in which the motive forceof gas compression is provided by high pressure steam generated withinthe plant by a steam turbine.
 7. The process according to claim 1 inwhich the motive force for gas compression is wholly or in part providedby the cooling system in the secont reactor zone.
 8. The processaccording to claim 1 wherein the temperature of the gaseous reactantscovering the first reactor zone are in the region of 180° C. to 220° C.